Method for operating catalytic reformers

ABSTRACT

A reforming process in which a hydrocarbon feed containing aliphatic hydrocarbons is converted to a hydrocarbon product comprising an increased proportion of aromatics by passage over a reforming catalyst in a sequence of moving bed reactors operating under reforming conditions including moderate hydrogen pressure. The process is applicable when a former fixed moving bed reformer has been converted to moving bed reactor operation with the recycle and other ancillary equipment retained so that moderate pressure (hydrogen partial pressure at least 11 barg) is required, usually with a catalysts such as Pt/Re which tend to exhibit excessive hydrogenolysis activity in moving bed service. The recycle hydrogen stream is split with a portion going to at least one reactor subsequent to the first reactor.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is related to U.S. patent applications Ser. Nos.10/690,801 (Publication No. U.S. 2004/0129605A1) and U.S. 60/564,133(Publication No. 2005/0274648), both of which relate to the conversionof fixed bed (semi-regenerative or cyclic) reforming units to operationwith moving bed reactors.

FIELD OF THE INVENTION

The invention relates generally to catalytic reformers. Moreparticularly, the invention relates to in improved method for operatinghigh pressure, fixed-bed catalytic reformer units which have beenconverted to units with continuous, moving-bed reactors.

BACKGROUND OF THE INVENTION

Catalytic reforming is an established refinery process used forimproving the octane quality of hydrocarbon feeds and, as noted in U.S.2004/0129605A1, reforming processes were traditionally operated assemiregenerative or cyclic processes using fixed bed reactors or, morerecently, as continuous processes using moving bed reactors. Proposalshave also been made for combining fixed and moving bed reactors with theregeneration mode being appropriate to the reactor types used in thehybrid configuration, so that the fixed bed reactors have retained thefixed bed type regeneration, usually semiregenerative, and the movingbed reactors in the unit have retained the dedicated moving bedregenerator. Units of this hybrid type are disclosed, for example, inU.S. 4,498,973; U.S. 5,190,638; U.S. 5,190,639; U.S. 5,5,196,110; U.S.5,5,211,838; U.S. 5,5,221,463; U.S. 5,354,451; U.S. 5,368,720 and U.S.5,417,843 as well as in the technical literature, for example, in NPRAPaper No. AM-96-50 “IFP Solutions for Revamping Catalytic ReformingUnits” (1996 NPRA Annual Meeting, Mar. 17-19, 1996) which describes amoving bed reforming unit in which two moving bed reactor stacks share acommon regenerator. UOP has recently announced its CycleX® Process forincreased hydrogen production from a fixed bed reforming unit by theaddition of a circulating catalyst reactor as the final reactor in thereactor sequence. This reactor is provided with its own heater andregenerator as an expansion of existing assets rather than as asubstitution of them. See NPRA Paper AM-03-93 and the UOP technical datasheet at http://www.uop.com/objects/cyclextechsheet.pdf.

In semiregenerative reforming, the entire reforming process unit isoperated by gradually and progressively increasing the temperature tocompensate for deactivation of the catalyst caused by coke deposition,until finally the entire unit is shut-down for regeneration andreactivation of the catalyst which is carried out with the catalystremaining in the reactor. In cyclic reforming, the reactors areindividually isolated by various piping arrangements. The catalyst isregenerated and then reactivated while the other reactors of the seriesremain on line. A “swing reactor” temporarily replaces the reactor whichis removed from the series for regeneration and reactivation of thecatalyst, which is then put back in the series. In continuous reforming,the reactors are moving-bed reactors with continuous or intermittentaddition and withdrawal of catalyst through which the catalyst movesprogressively before it is passed to a regeneration zone forregeneration and rejuvenation before being returned once again to thereactor. In the regenerator, at least a portion of the deposited coke isburned off and the regenerated catalyst is recycled to the reactor tocontinue the cycle. Commercial continuous reforming units may have thereactors arranged in a side-by-side or in a stacked configuration.Because the continuous mode of operation with its frequent regenerationcan tolerate a higher rate of coke lay-down on the catalyst, it ispossible to operate continuous units at lower and more favorablepressures than those normally used with semi-regenerative and cyclicunits in which it is important or at least desirable to extend catalystlife between successive regenerations.

Semiregenerative and cyclic reforming units may be converted tocontinuous moving-bed units to take advantage of the improved yield ofhigher octane reformate and hydrogen associated with continuous lowerpressure operation but the conversions which have so far been consideredare essentially entire unit replacements which require replacement ofall existing vessels and most of the ancillary equipment as well asinstallation of an integrated catalyst regenerator which is one of themost costly items in the conversion. The cost of the regenerator can beas much as about seventy percent of the total cost required for theconversion. U.S. patent applications Ser. Nos. 10/690,081 (PublicationNo. U.S. 2004/0129605A1) and 60/564,133 describe different conversiontechniques by which fixed bed reformers may be converted to moving bedoperation without the major expense normally associated with a completeconversion. The technique described in U.S. patent application Ser. No.10/690,081 replaces the fixed bed reactors with moving bed reactors butretains the existing heaters and piping and eliminates the need for anindividual, dedicated regenerator either by utilizing off-siteregeneration or by utilizing a regenerator of another unit. Thetechnique described in U.S. patent application Ser. No. 60/564,133, bycontrast, enables a fixed bed unit to be converted to a unit with movingbed reactors while converting existing fixed-bed reactor vessels to useas regenerator vessels which are switched in a cyclic sequence betweenfilling, regeneration/rejuvenation and emptying modes.

As noted above, a major advantage of moving bed operation is that it maybe operated under a regime in which hydrogen production and conversionof aliphatics to aromatics proceeds to a more favorable equilibriumunder the prevailing lower pressure regime. While catalyst cokingproceeds at a greater rate under these conditions of reduced hydrogenpartial pressure, this is acceptable when the catalyst is regeneratedafter only a relatively short cycle time in the reactor stack. Theconditions normally encountered in continuous, moving bed reforming do,however, impose their own limitations on the choice of catalyst. Whilefixed bed reformers, whether of the cyclic or semi-regenerative typetraditionally used catalysts based on the platinum-rhenium (Pt/Re)combination, these catalysts, even with additional metallic promoterssuch as iridium, were not optimal for moving bed operation. Under thelow pressure conditions used in conventional moving bed reformers,typically below 11 bar (gauge), the normal commercially availablecatalysts are platinum-tin (Pt/Sn) and this combination of metals isnotably more active than the Pt/Re combination for dehydrocyclization atlow hydrogen partial pressures. A problem arises, however, with the lowcost revamps described in the two applications referred to above, whichare intended to provide a route to moving bed reactor operation whileavoiding the major expense of regenerator acquisition and replacement ofexisting ancillary equipment including, especially, the recycle gascompressors and furnaces which are very expensive. The retention of therecycle circuit often limits the magnitude of the pressure reductionwhich can be achieved during the revamp. If the minimum revampedpressure has to be kept above 11 bar (approx. 160 psig), typicallybetween about 13 and 35 barg (approx. 190 to 510 psig), catalysts basedon the Pt/Re combination used at the higher pressures of traditionalfixed bed operation may be preferred.

Existing commercial practice in fixed bed reactors using Pt/Recatalysts, enables sufficient coke to be rapidly deposited, even nearthe inlet of the first reactor, to mitigate the undesirablehydrogenolysis activity associated with the rhenium; in moving bedoperation, by contrast, the catalyst closest to the inlet of the firstreactor, will consistently have a coke level which is substantiallylower than that found in fixed bed operation with Pt/Re catalysts, evenshortly after start of run since fresh, uncoked catalyst is continuouslyadded to the inlet of the first reactor. The freshly added Pt/Recatalyst in the first moving bed reactor does not acquire sufficientcoke sufficiently quickly to mitigate the inherent hydrogenolysisactivity to an acceptable level within a reasonable period of time.Moving bed pilot plant studies with Pt/Re catalyst have shown that evenat relatively slow catalyst circulation rates, the steady state cokelevel anywhere in the lead reactor would be well below 1 weight percenton catalyst. While some of the hydrogenolysis can be mitigated bypresulfiding the freshly regenerated Pt/Re catalyst before it is addedto the top of the lead reactor, this sulfur is stripped from thecatalyst relatively quickly and does not sufficiently suppresshydrogenolysis as the catalyst moves slowly down through the leadreactor. As a result, the refiner seeking to avail himself of a low costfixed-to-moving bed conversion faces a dilemma: if existing recycle andother equipment is retained, requiring operation at moderate pressure,it may not be economic to use the catalysts which are preferred atmoderate pressure conditions because of excessive hydrogenolysis; on theother hand, the Pt/Sn catalysts used in commercial moving bed reformersare not preferred at the higher pressures. Thus, there is a need forresolving this dilemma in a way which permits operation of a revampedunit with the Pt/Re and other catalysts which afford the bestperformance in the moderate pressure regime imposed by the equipmentlimitations.

SUMMARY OF THE INVENTION

The performance of catalysts in moving bed reactors, especially of Pt/Reand Pt/Ir based catalysts, at moderate pressures can be improved bysplitting the recycle gas stream, such that only a portion of the totalrecycle gas is returned to the inlet of the first reactor. Thisdiversion of hydrogen-rich recycle gas away from the first reactor, willdecrease the H2/hydrocarbon molar ratio in the first reactor, andincrease the catalyst coke level within the first reactor to a level atwhich the undesirable hydrogenolysis reactions are suppressed, Although,it is desirable to minimize the catalyst coke level in most commercialreforming processes, in this unique case, the intentional cokedeposition in the lead moving bed reactor(s) will increase hydrogenpurity, hydrogen yield, reformate yield, and aromatics yield. Areduction of the total recycle rate of hydrogen-rich gas would, bycontrast, not produce the same benefit as this diversion of a portion ofthe recycle gas flow away from the first reactor: if the total recyclegas flow rate is too low, it will cause an undesirable increase in therate of coke formation across all reactors, including those in which thecatalyst already has sufficient coke to inhibit rhenium hydrogenolysisactivity. The splitting of the hydrogen recycle among the reactors,therefore, reduces undesired hydrogenolysis in the precise locationwhere, under these conditions, it has become a problem. In addition, thediversion of the recycle gas also has the advantage of reducing thepressure drop in the recycle circuit so as to permit lower pressureoperation of the recycle gas compressor or a higher relative naphthafeed rate.

According to the present invention, therefore, the reforming process iscarried out in a sequence of moving bed reactors operating underreforming conditions including a hydrogen partial pressure of at least11 bar gauge (approx, 160 psig), has the recycle hydrogen resulting fromthe reforming reactions recycled to the first reactor and to at leastone reactor in the sequence subsequent to the first reactor,

THE DRAWING

FIG. 1 shows a continuous moving-bed reforming unit built from anexisting high pressure fixed-bed reformer unit with a split hydrogenrecycle circuit.

DETAILED DESCRIPTION

The present invention is applicable to catalytic naphtha reforming, thatis to the process in which a hydrocarbon feed in the naphtha boilingrange is subjected to reactions at elevated temperature includingdehydrogenation, dehydrocyclization, isomerization and hydrocracking toconvert aliphatic hydrocarbons in the naphtha feed to aromatics so as toresult in a product comprising an increased proportion of aromatics(relative to the feed). Depending on the properties of the naphthafeedstock (as measured by the paraffin, olefin, naphthene, and aromaticcontent) and catalysts used, the reformate product can be produced withvery high concentrations of toluene, benzene, xylene, and otheraromatics useful in gasoline blending and petrochemical processing.Hydrogen, a significant by-product, is separated from the reformate forrecycling and use in other refinery processes. While the reactionsinvolved in the overall reforming process include both exothermic andendothermic components, the overall reaction is endothermic and requiressubstantial amounts of process heat to carry it to the desired point.The older type of fixed bed reformers typically operated at moderate tohigh hydrogen pressures in order to extend the cycle life of thecatalyst between regeneration cycles; the more modem continuouscatalytic reformers, however, are capable of operating in a lower, morefavorable pressure regime. The present invention is principallyapplicable to operation with moving bed reactors in the high to moderatepressure range, typically above about 11 barg and usually above about 12barg. Pressure of this order, typically in the range of 15-27 barg, willnormally be encountered in moving bed reactors in units made by theconversion of older, fixed bed units where the compressor, furnaces, andassociated recycle equipment have been retained, for example, asdescribed in applications Ser. Nos. 10/690,081 and 60/564.433, forreasons of economy.

FIG. 1 depicts a catalytic reformer unit which has been converted tomoving bed reactor operation from fixed bed (cyclic orsemi-regenerative) operation as described in U.S. patent applicationSer. No. 10/690,081 using offsite or community catalyst regeneration. Inthis case, recycle gas diverted away from the first reactor inlet ispassed through a low-pressure drop heat exchanger, and then combinedwith the first reactor effluent. The combined stream, containing boththe diverted recycle gas and effluent from the first reactor, thenpasses through the first reheat furnace before entering the inlet of thesecond reactor. The recycle gas diverted away from the first reactorinlet, is passed through a dedicated, low-pressure drop, feed/effluentheat exchanger to minimize circuit pressure drop. This strategy allowsthe maximum reduction in the pressure drop across both the primaryfeed/effluent heat exchanger train, the preheat furnace, and the firstreactor, due to a decrease in the recycle gas flow rate through each.

In the unit shown in FIG. 1, the hydrocarbon feed enters through line 10and is combined with recycle hydrogen from line 11. The combinedhydrocarbon/hydrogen feed then passes through heat exchanger 12 in whichit picks up heat from the effluent from the final reactor (three shown).The combined feed then passes through furnace 13 to bring it to therequired temperature to enter first reactor 15 in which the reformingreactions commence. Reactor 15, like the second and third stage reactors16 and 17, is a moving bed reactor that replaces the former fixed bedreactor from a cyclic or semi-regenerative reforming unit. Normally, thereactor vessel will be a radial flow reactor to minimize pressure drop.Additional hydrogen recycle gas from line 20 which has picked up heatfrom the effluent from third stage reactor 17 in heat exchanger 21 joinsthe effluent from first stage reactor 15 and the combinedhydrocarbon/hydrogen stream then passes to the second stage reactor 16by way of second furnace 22 and from second stage reactor 16 in theconventional manner to third stage reactor 17 by way of third furnace23. The effluent from the third stage reactor is split and passesthrough heat exchangers 12 and 21 to heat the recycle streams, with theratio of hot effluent between the two exchangers being controlled byvalve 24 in accordance with the selected recycle split ratio. The spliteffluents are then recombined in line 26 before passing to air cooler30, separator 31 from which the reformate product is removed throughline 32 with recycle hydrogen passing to recycle compressor 33. Therecycle hydrogen is split between lines 35 and 36 to be directed,respectively, to the first stage feed through line 11 and to the secondstage feed through line 20 with the split ratio being controlled bymeans of valve 37. If desired, the recycle hydrogen may be split betweenall three reactors using line 38 for this purpose although, given thefact that the majority of the dehydrocyclization takes place in thethird stage reactor this will not normally be desired if conditionsconducive to most favorable equilibrium are to be attained.

For the purposes of this illustration, the reactors are shown as beingoriented side by side but either side by side or vertically stackedconfigurations may be used. Similarly, three reactors are shown,arranged in series. The split recycle arrangement is, however,applicable also to cases with two reactors in series and cases with morethan three reactors in series. In each case, the objective is to reducethe proportion of recycle gas to the point where sufficient cokeaccumulation on the catalyst will take place to suppress or inhibit theundesired hydrogenolysis reactions in the lead reactors. Because thesereactions take place for the most part over the freshly regeneratedcatalyst, the hydrogen partial pressure at the first stage reactor inletis the most sensitive variable and the split ratio should be adjustedaccordingly; the split, if any is provided between or among theremaining reactors is less significant in its effects since by the timethe catalysts enters the subsequent stages, coke deposition will havetaken place and hydrogenolysis reduced. Thus, in most case, a splitbetween the first and second stages will be all that is required, asshown in FIG. 1, where all the recycle gas diverted away from the firstreactor, is combined with the effluent from the first reactor.

FIG. 1 depicts a configuration in which the recycle gas diverted fromthe first reactor is passed through a separate feed/effluent heatexchanger (21) prior to being combined with the first reactor effluent.The primary reason for sending the split recycle gas streams throughseparate feed effluent heat exchangers is that in most reformers thefeed, along with all the recycle gas, is preheated in the existingfeed-effluent heat exchanger. When the split recycle is used, however,it may be desirable to heat the recycle gas stream that goes to theinlet of the second reactor separately to avoid bypassing some of thefeed to the inlet of the second reactor. This may not be necessary,however, if the reheat furnace upstream of each subsequent reactor has,as is normally the case, sufficient capacity to heat the feed and thesplit recycle gas stream to the required temperature for the reactor. Insuch cases, the recycle gas diverted away from the first reactor neednot be passed through a separate feed/effluent heat exchanger prior tobeing redistributed between the inlets to any of the subsequentreactors.

The split recycle mode of operation has two key benefits:

It allows lower circuit pressure drop across the existing feed effluentheat exchangers, preheat furnace, and lead reactor(s), than would bepossible if all recycle gas were returned to the inlet of the firstreactor. This is important in the revamp of an existing fixed bed unitwhere the existing recycle gas compressor could otherwise limit theextent of the revamp. The reduced pressure drop across the unit providesfor (a) lower pressure operation of the existing recycle gas compressorat constant naphtha feed rate, (b) higher naphtha feed rate at constantpressure or (c) some combination of (a) and (b). Second, the loweroperating pressure also improves yields of reformate and hydrogen fromthe reformer.

The lower H₂/hydrocarbon molar ratio, in the lead reactor(s), willincrease the catalyst coke level. Nominally, two or more weight percentcoke is desirable on Pt/Re catalysts to reduce undesirablehydrogenolysis activity associated with the rhenium. This intentionalcoke lay down in the first and possibly second moving bed reactors, willincrease hydrogen purity, hydrogen yield, reformate yield, and aromaticsyield.

Reforming process simulations indicate that H₂ production, H₂ purity andreformate yield will increase by a relative 25%, 1.2% and 0.35%,respectively, for each 69 kPag (10 psig) reduction in pressure(2400-2070 kPag (350-300 psig) range). The pilot unit data shown Table 1summarizes the yield benefits of increasing the lead reactor coke level,for moving bed Pt/Re catalysts. The data in Table 1 was collected at thefollowing process conditions: Naphtha feed rate=1.5 hr⁻¹ WHSV, reactorpressure=2400 Kpag (350 psig), H₂/hydrocarbon molar ratio=1.8, reformateoctane=98 C5+RON.

The ratio between the volume of hydrogen going to the first stagereactor and subsequent reactors may be determined empirically, dependingupon the catalyst being used and the conditions employed in the firststage reactor. Normally, it may be expected that the fraction of thetotal recycle gas going to the first stage reactor could be as little as25-50%.

The conversion of the fixed bed unit to moving bed reactor operation maysuitably be carried out as described in application Ser. No. 10/690,801or 60/564,133, to which reference is made for a description of theconversion technique. The split recycle mode is however applicable touse with units of other types with a sequence of at least two moving bedreactors operating at moderate to high pressures. In the conversion offixed bed units, the fixed-bed reactor vessels will be converted to amoving bed reactor unit which allows continuous or intermittent additionof fresh or regenerated catalyst to its catalyst inlet and continuous orintermittent removal of spent catalyst from the catalyst outlet of theunit after the final reactor in the unit. The unit will be provided withcatalyst feed facilities for continuously or intermittently chargingfresh or regenerated catalyst in a continuous or intermittent mode ofoperation to the moving-bed reactor. In addition, spent catalystrecovery facilities will be added for collecting the spent catalyst,storing it temporarily, and transferring it to a catalyst regenerationfacility. The moving-bed reactors, the catalyst feeding facilities andthe catalyst recovery facilities are operatively connected betweenthemselves and to the existing facilities (piping, ancillary equipment)of the fixed-bed unit that will not require replacement. In addition,the recycle gas circuit will be modified to split the stream between thefirst reactor and at least one subsequent reactor, normally between thefirst reactor and the second reactor, as described above.

The moving-bed reactors are operated at an effective pressure to improvereformate quality and yield compared to the quality and yield from thefixed-bed unit before the conversion. The moving-bed reformer reactorsof the converted unit may be operated at an effective pressure that issufficiently low to improve substantially the reformate quality andyield as compared to the reformate quality and yield obtained from thefixed-bed unit before conversion. The pressure is, however, maintainedduring normal operation at a value which is sufficiently high to allowthe use of the existing equipment of the fixed-bed catalytic reformerunit including the recycle compressor, heat exchangers and furnaces. Theuse of a higher pressure than typical for a fully integrated continuousreactor-regenerator is desirable in that it enables the rate of catalystflow for regeneration to be reduced (relative to that of an integratedunit) and so relieves the burden of catalyst handling while operationwith the split recycle permits the use of the Pt/Re based catalystspreferred in the higher pressure operation at pressures above about 11barg and usually above 12 barg. As noted above, pressures in the range15 to 30 barg, e.g. 24-27 barg may be expected with these convertedunits.

The catalyst may be any reforming catalyst which may be found suitablefor the operation of the unit. Invariably, these will be platinum basedbimetallic or trimetallic catalysts with the platinum combined with asecond metal, normally tin, rhenium or iridium. The Pt/Re metal systemswhich have conventionally been used in fixed bed units will normally besuitable at the pressures encountered in the converted units with theobvious proviso that the catalyst must be fabricated into a formsuitable for moving bed operation, i.e. in bead form, typically with abead size of 1 to 10, e.g. 1-3 mm, so as to minimize attrition. Ifdesired and operating conditions permit, the Pt/Sn system may be used,as is typical of continuous catalytic reforming (CCR) units. The optimalchoice of catalyst may be made empirically depending upon unitconfiguration and operating conditions, with a wide range of catalystsbeing commercially available. As noted above, however, the presentinvention is of particular utility with the Pt/Re catalysts which,although favored for fixed bed application, tend to exhibit excessivehydrogenolysis activity in moving bed operation. It is also applicable,however, to other catalyst metals and metal systems in which excessivehydrogenolysis is encountered when the catalyst is used in moving bedoperation. This would include, for example, bimetallic Pl/Ir catalysts.Whichever metal system is used, however, third modifying metals may alsobe present, for example, iridium, rhodium or ruthenium, e.g.trimetallics such as Pt/Re/Ir. The amount of the platinum group metalwill typically be from 0.01 to 5 wt. pct, more usually from 0.1 to 0.3wt pct of the metal of which most will be platinum. The support materialwill conventionally be alumina. The catalyst preferably has a highsurface area, for example, from 100 to 200 m²g⁻¹. As is conventional,halogenation treatment will be applied in order to maintain the acidfunction of the catalysts and conventional materials and techniques willbe directly applicable. Descriptions of typical reforming catalysts andhalogenation treatments are to be widely found in the technical andpatent literature.

Regeneration of the catalyst may take place as described in applicationSer. No. 10/609,801, that is, using an offsite catalyst regenerationfacility or, alternatively, a community onsite continuous catalystregeneration facility, i.e., a continuous catalyst regenerator that isshared between more than one reactors or a non-continuous onsiteregeneration facility for one or more catalytic reformer units.

The composition of hydrocarbon feeds and products will be dictated byrefinery equipment and conditions and product needs. In brief, the feedwill typically be a petroleum naphtha with a boiling range of 20°-250°C. (about 70°-480°F.), more usually 60°-200° C.(about 140°-390° F.)although higher end points are frequently encountered. Typical feedsinclude straight run naphthas, cracked naphthas, synthetic naphthas suchas shale oil naphtha, hydrocracked naphthas and blends of the above.Feed pretreatment to reduce contaminants especially sulfur and nitrogento very low levels typical of reforming is required in the conventionalmanner. Conditions for the moving bed reactors will typically includetemperatures form 425°-650° C. (about 800° to 1200° F.) usually425°-550° C. (about 800°-1000° F.) at the pressures described above,with other parameters such a space velocity and overallhydrogen/hydrocarbon ratio conventional for moving bed operation.Descriptions of typical reforming processes are to be widely found inthe technical and patent literature.

1. A reforming process in which a hydrocarbon naphtha feed containingaliphatic hydrocarbons is converted to a hydrocarbon product comprisingan increased proportion of aromatics by passage over a reformingcatalyst in a sequence of moving bed reactors operating under reformingconditions including a hydrogen partial pressure of at least 11 barg inwhich hydrogen is fed to at least one reactor subsequent to the firstreactor in addition to the first reactor of the sequence.
 2. A processaccording to Claim 1 in which the hydrogen fed to the reactors isrecycle hydrogen separated from the reforming products.
 3. A processaccording to Claim 2 in which the sequence of moving bed reactorsincludes three reactors and the hydrogen is fed to the first reactor andthe second reactor of the three-reactor sequence.
 4. A process accordingto Claim 2 in which the hydrogen separated from the reforming productsis heated by heat exchange with the reforming products prior to enteringthe first reactor.
 5. A process according to Claim 3 in which thehydrogen separated from the reforming products is heated by heatexchange with the reforming products prior to entering the first and thesecond reactor of the sequence.
 6. A process according to Claim 5 inwhich the hydrogen separated from the reforming products is split intotwo streams which are heated separately by heat exchange with thereforming products.
 7. A process according to Claim 1 in which thereforming catalyst is a Pt/Re catalyst.
 8. A process according to Claim1 which is operated at a pressure of 12 to 30 barg measured at the inletof the first reactor.
 9. A process according to Claim 8 which isoperated at a pressure of 12 to 27 barg measured at the inlet of thefirst reactor.
 10. A process according to Claim 9 in which the reformingcatalyst is a Pt/Re catalyst.
 11. A reforming process in which ahydrocarbon naphtha feed in the boiling range of 60° to 200° C. whichcontains aliphatic hydrocarbons is converted to hydrogen and ahydrocarbon product comprising an increased proportion of aromaticsrelative to the naphtha feed by passage over a reforming catalyst underreforming conditions including a temperature in the range of 425°-650°C. and a hydrogen partial pressure of at least 11 barg in a sequence ofmoving bed reforming reactors following which hydrogen is separated fromthe hydrocarbon reforming product and recycled to the first reactor ofthe sequence and to at least one reactor in the sequence subsequent tothe first reactor.
 12. A process according to Claim 11 in which thesequence of moving bed reactors includes three reactors and the hydrogenis fed only to the first reactor and the second reactor of thethree-reactor sequence.
 13. A process according to Claim 12 in which thehydrogen separated from the reforming products is heated by heatexchange with the reforming products prior to entering the firstreactor.
 14. A process according to Claim 13 in which the hydrogenseparated from the reforming products is heated by heat exchange withthe reforming products prior to entering the first and the secondreactor of the sequence.
 15. A process according to Claim 14 in whichthe hydrogen separated from the reforming products is split into twostreams which are heated separately by heat exchange with the reformingproducts.
 16. A process according to Claim 11 in which the reformingcatalyst is a Pt/Re catalyst.
 17. A process according to Claim 11 whichis operated at a pressure of 12to 30 barg measured at the inlet of thefirst reactor.
 18. A process according to Claim 17 which is operated ata pressure of 12to 27 barg measured at the inlet of the first reactor.19. A process according to Claim 18 in which the reforming catalyst is aPt/Re catalyst.
 20. A reforming process in which a hydrocarbon naphthafeed in the boiling range of 60° to 200° C. which contains aliphatichydrocarbons is converted to hydrogen and a hydrocarbon productcomprising an increased proportion of aromatics relative to the naphthafeed by passage over a reforming catalyst comprising platinum andrhenium on an alumina support under reforming conditions including atemperature in the range of 425°-550° C. and a hydrogen partial pressureof at least 12 barg in a sequence of three moving bed reforming reactorsfollowing which hydrogen is separated from the hydrocarbon reformingproduct and split into two streams, one of which is heated by heatexchange with reforming product effluent from the third reactor of thesequence before being recycled to the first reactor of the sequence withthe other stream being separately heated by heat exchange with reformingproduct effluent from the third reactor of the sequence before beingrecycled to the second reactor of the sequence.